Slurry phase polymerisation process

ABSTRACT

Process for producing a multimodal polyethylene in at least two reactors connected in series, in which 20-80 wt % of a high molecular weight (HMW) polymer is made in suspension in a first reactor and 20-80 wt % of a low molecular weight (LMW) polymer is made in suspension in a second reactor. The ratio of the average activity in the LMW reactor to the average activity in the HMW reactor is from 0.25 and 1.5, where average activity in each reactor is defined as the rate of polyethylene produced in the reactor (kgPE/hr)/[ethylene concentration in the reactor (mol %)×residence time in the reactor (hours)×feed rate of catalyst into the reactor (g/hr)], residence time being defined as the mass of the polymer in the reactor (kg)/the output rate of polymer from the reactor (kg/hr). The volume of the second reactor is at least 10% greater than the volume of the first reactor, and the ratio of length to diameter of the first reactor, L/D(1), is greater than that of the second reactor, L/D(2).

This application is the U.S. national phase of International ApplicationNo. PCT/EP2008/061366 filed 29 Aug. 2008 which designated the U.S. andclaims priority to European Application No. 07253487.8 filed 3 Sep.2007, the entire contents of each of which are hereby incorporated byreference.

The present invention is concerned with olefin polymerisation in slurryphase reactors, and more particularly with polymerisation in two or morereactors arranged in series.

Slurry phase polymerisation of olefins is well known wherein an olefinmonomer and optionally olefin comonomer are polymerised in the presenceof a catalyst in a diluent in which the solid polymer product issuspended and transported.

Polymerisation is typically carried out at temperatures in the range50-125° C. and at pressures in the range 1-100 bara. The catalyst usedcan be any catalyst typically used for olefin polymerisation such aschromium oxide, Ziegler-Natta or metallocene-type catalysts.

Many multiple reactor systems employ loop reactors, which are of acontinuous tubular construction comprising at least two, for examplefour, vertical sections and at least two, for example four horizontalsections. The heat of polymerisation is typically removed using indirectexchange with a cooling medium, preferably water, in jackets surroundingat least part of the tubular loop reactor. The volume of each loopreactor of a multiple reactor system can vary but is typically in therange 10-200 m³, more typically 50-120 m³. The loop reactors employed inthe present invention are of this generic type.

Typically, in the slurry polymerisation process of polyethylene forexample, the slurry in the reactor will comprise the particulatepolymer, the hydrocarbon diluent(s), (co) monomer(s), catalyst, chainterminators such as hydrogen and other reactor additives In particularthe slurry will comprise 20-75, preferably 30-70 weight percent (basedon the total weight of the slurry) of particulate polymer and 80-25,preferably 70-30 weight percent (based on the total weight of theslurry) of suspending medium, where the suspending medium is the sum ofall the fluid components in the reactor and will comprise the diluent,olefin monomer and any additives; the diluent can be an inert diluent orit can be a reactive diluent in particular a liquid olefin monomer;where the principal diluent is an inert diluent the olefin monomer willtypically comprise 2-20, preferably 4-10 weight percent of the slurry.

The slurry is pumped around the relatively smooth path endless loopreaction system at fluid velocities sufficient to maintain the polymerin suspension in the slurry and to maintain acceptable cross-sectionalconcentration and solids loading gradients. Slurry is withdrawn from thepolymerisation reactor containing the polymer together with the reagentsand inert hydrocarbons, all of which mainly comprise inert diluent andunreacted monomer. The product slurry comprising polymer and diluent,and in most cases catalyst, olefin monomer and comonomer can bedischarged intermittently or continuously, optionally usingconcentrating devices such as hydrocyclones or settling legs to minimisethe quantity of fluids withdrawn with the polymer.

In multiple reactor polymerisations, the composition of the slurrywithdrawn from the final reactor depends on many factors apart from thecomposition of the product actually polymerised in the final reactor: italso depends on the desired final product, and the reaction conditionsand relative proportions of products in any upstream reactors. Thereaction conditions required in the final reactor are also affected bythe reaction conditions in upstream reactors, particularly the impact ofcatalyst productivity in upstream reactors on the average activitypotential under downstream reaction conditions. Thus the control of theslurry composition withdrawn from the final reactor and also the processconditions associated therewith is more complex than in the case of asingle reactor.

One issue which can affect all of the above factors is the relative sizeof the two reactors. There are many conflicting requirements affectingthe optimisation of the volume and dimensions of the two reactors. In amultiple reactor polymerisation, the second and any subsequent reactorsneed to be large enough to handle not only the polymer produced in thatreactor, but also the polymer transferred from the previous reactor orreactors. This would imply that the second and subsequent reactorsshould be larger than the upstream reactors in order maintain similarspace time yields. EP 057420A discloses a a two-reactor system in whichthe second reactor is twice the volume of the first one. We have foundthat one disadvantage of this arrangement is that the heat removalrequirement, often a production constraint, of the larger downstreamreactors is greater than that of the upstream reactor. Accordingly it isnot obvious what the optimum size ratio of the reactors should be. Inparticular, where a reactor system is designed to operate differentcatalyst types (eg Ziegler-Natta, chromium and/or metallocene), or acatalyst system where the average activity or production ratio requiredvaries significantly between reactors under different operating regimes,the ideal ratio of sizes for the reactors in each case is likely to bedifferent, making it difficult to select an ideal size profile. Theactivity profile of under constant reaction conditions also variessignificantly between Ziegler-Natta, chromium, metallocene and/or latetransition metal catalyst systems.

However we have found that the most optimum reactor size ratio is one inwhich the second reactor is at least 10 vol % larger than the firstreactor, and additionally the ratio of length to diameter of the firstreactor is greater than that of the second reactor.

Thus in its first aspect, the present invention provides a for producinga multimodal polyethylene in at least two reactors connected in series,in which 20-80 wt % of a high molecular weight (HMW) polymer is made insuspension in a first reactor and 20-80 wt % of a low molecular weight(LMW) polymer is made in suspension in a second reactor, wherein theratio of the average activity in the LMW reactor to the average activityin the HMW reactor is from 0.25 and 1.5, where average activity in eachreactor is defined as the rate of polyethylene produced in the reactor(kgPE/hr)/[ethylene concentration in the reactor (mol %)×residence timein the reactor (hours)×feed rate of catalyst into the reactor (g/hr)],residence time being defined as the mass of the polymer in the reactor(kg)/the output rate of polymer from the reactor (kg/hr), and whereinthe volume of the second reactor is at least 10%, preferably at least30% and more preferably at least 50% greater than the volume of thefirst reactor, and the ratio of length to diameter of the first reactor,L/D(1), is greater than that of the second reactor, L/D(2), andpreferably at least 20% greater.

A reactor system in which the volume of the second reactor is at least10% greater than the volume of the first reactor enables the totalreactor volume to be minimised whilst providing sufficient flexibilityto handle different operating conditions and catalysts. It isparticularly advantageous in cases where the catalyst activity in thefirst HMW reactor is high, as the larger second reactor size enableshigher residence times to be employed for a given block ratio. We havefound that it is possible to overcome the difficulties of differentrequirements for heat removal in the two reactors by ensuring that theratio of length to diameter of the first reactor, L/D(1), is greaterthan that of the second reactor, L/D(2), preferably at least 20%greater, and most preferably at least 30% greater. Typically, the ratioof L/D(1) to L/D(2) is greater than 1.5, most preferably greater than 2.An increased L/D provides a greater surface area per unit volume, whichenables a faster rate of heat removal, as the ability to cool a reactordepends on the surface area available to which cooling can be applied.Thus if the cooling requirements of the two reactors are the same, thelarger LMW reactor can have a lower L/D than the smaller HMW reactor.Thus the invention enables the heat transfer capability of each reactorto be balanced whilst also minimising total reactor volume.

Generally it is preferred that the ratio of length to diameter (L/D) ofthe first HMW reactor is greater than 500, preferably between 750 and3000, and most preferably greater than 800, for example 800-1500.Generally it is preferred that the ratio of length to diameter (L/D) ofthe second LMW reactor is greater than 200, preferably 200-1000, andmost preferably 250-750, for example 300-550.

Usually each of the reactors has an internal volume greater than 10 m³,more commonly greater than 25 m³ and in particular greater than 50 m³.Typical ranges are 75-200 m³ and more particularly 100-175 m³.

Average activity in each reactor is defined as the rate of polyethyleneproduced in the reactor (kgPE/hr)/[ethylene concentration in thereactor(mol %)×residence time in the reactor(hours)×feed rate ofcatalyst into the reactor (g/hr)], residence time being defined as themass of the polymer in the reactor (kg)/the output rate of polymer fromthe reactor (kg/hr). If no additional catalyst is added to the secondreactor, when calculating the ratio of average activities the flow rateof catalyst in the two reactors is considered to be the same. Ifadditional catalyst is added to the second reactor, the flow rate intothe second reactor is considered to be the sum of the flowrate ofcatalyst from the first reactor plus the flowrate of additional freshcatalyst added directly into the second reactor. Alternatively, activityin each reactor may be calculated based on catalyst residues in thepolymer produced in each reactor, as is well known, and the activityratio calculated from this.

The residence time is defined as the mass of the polymer in the reactor(kg)/the output rate of polymer from the reactor (kg/hr). In a casewhere polymer is recycled back into the reactor, for example when ahydrocyclone is employed downstream of the reactor, the output rate ofpolymer is the net output rate (ie polymer withdrawn less polymerrecycled).

Preferably the multimodal polyethylene has a shear ratio of at least 15,generally between 15 and 50, and preferably between 21 and 35. By “shearratio” is the ratio of the high load melt index HLMI of the polyethyleneto the MI₅ of the polyethylene. The HLMI and MI₅ are measured accordingto ISO Standard 1133 at a temperature of 190° C. using loads of 21.6 kgand 5 kg respectively. MI₂ is similarly measured but using a load of2.16 kg.

The HLMI of the multimodal polyethylene exiting the second reactor ispreferably between 1 and 100 g/10 min, and more preferably between 1 and40 g/10 min.

In one embodiment the catalyst employed for the polymerisation is aZiegler-Natta catalyst. In this case, it is preferred that the ratio ofLMW to HMW polymer is from 40:60 to 60:40.

In multiple reactor polymerisations, the composition of the slurrywithdrawn from the final reactor depends on many factors apart from thecomposition of the product actually polymerised in the final reactor: italso depends on the desired final product, and the reaction conditionsand relative proportions of products in any upstream reactors. Thereaction conditions required in the final reactor are also affected bythe reaction conditions in upstream reactors, particularly the impact ofcatalyst productivity in upstream reactors on the average activitypotential under downstream reaction conditions. It is generallydesirable that the majority of the liquid components withdrawn with thepolymer from the final reactor are separated in a flash tank at atemperature and pressure such that they can be recondensed just bycooling, without recompression. The remaining liquid components notremoved by this process are separated in a second flash tank operatingat a lower pressure, and these need to be recompressed in order to berecycled. The advantage of this process, which is referred tohereinafter as a “medium pressure flash” process, is that only a smallproportion of the vaporised liquid components need to be recompressed inorder to be recondensed. We have found that by careful control of thereaction conditions it is possible to ensure that a “medium pressureflash” process is able to be operated without the need for recompressionof the liquid vaporised in the first flash tank.

It is preferred that the ratio of solids concentration in the firstreactor to that in the second reactor is maintained at less than 1.0,preferably between 0.6 and 0.8, as this also assists in maintaining thebalance of average activity between the two reactors within the desiredrange. The solids concentration is the average weight of polymerrelative to the total weight of the slurry.

Generally the solids concentration in the second LMW reactor is at least35 wt %, most preferably between 45 wt % and 60 wt %. The solidsconcentration in the HMW reactor is usually between 20 wt % and 50 wt %,more preferably between 25 wt % and 35 wt %. In this case it ispreferred to concentrate the solids transferred from the first reactorto the second reactor using a settling zone and/or hydrocyclone. Acomonomer-free diluent stream may be introduced upstream of thehydrocyclone to reduce the proportion of comonomer transferred to thedownstream reactor, thus increasing the density of the polymer producedin the LMW reactor.

A further aspect of the invention provides a process for producing amultimodal polyethylene in at least two reactors connected in series, inwhich 20-80 wt % of a high molecular weight (HMW) polymer is made insuspension in a first reactor and 20-80 wt % of a low molecular weight(LMW) polymer is made in suspension in a second reactor, in which thesolids concentration in the second LMW reactor, defined as the mass ofpolymer divided by the total mass of slurry, is at least 35 wt %, mostpreferably between 45 wt % and 60 wt %, and/or the the ratio of solidsconcentration in the first reactor to that in the second reactor ismaintained at less than 1.0, preferably between 0.6 and 0.8, and thevolume of the second reactor is at least 10%, preferably at least 30%and more preferably at least 50% greater than the volume of the firstreactor, and the ratio of length to diameter of the first reactor,L/D(1), is greater than that of the second reactor, L/D(2), andpreferably at least 20% greater. In this further embodiment it ispreferred that the ratio of the average activity in the LMW reactor tothe average activity in the HMW reactor is from 0.25 and 1.5, whereaverage activity in each reactor is defined as the rate of polyethyleneproduced in the reactor (kgPE/hr)/[ethylene concentration in the reactor(mol %)×residence time in the reactor (hours)×feed rate of catalyst intothe reactor (g/hr)], residence time being defined as the mass of thepolymer in the reactor (kg)/the output rate of polymer from the reactor(kg/hr). It is also preferred that the solids concentration in the HMWreactor is between 20 wt % and 50 wt %, more preferably between 25 wt %and 35 wt %.

We have found that maximising the solids concentration in the second LMWreactor relative to that in the first HMW reactor is an effective way ofincreasing the residence time in the second reactor relative to that inthe first, in order to balance the production rate in the two reactors.

In a preferred version of the invention, a slurry containing themultimodal polyethylene is transferred from the second of the tworeactors to a flash tank operating at a pressure and temperature suchthat at least 50 mol %, preferably that at least 80 mol %, morepreferably 90 mol %, most preferably 95 mol % of the liquid component ofthe slurry is withdrawn from the flash tank as a vapour. In thisembodiment it is preferred that the concentration in the flash tank ofcomponents having a molecular weight below 50, C_(lights), satisfies theequation C_(lights)<7+0.07(40−T_(c))+4.4(P_(c)+0.8)−7(C_(H2)/C_(Et))where T_(c) and P_(c) are respectively the temperature (in C) andpressure (MPa g) at the location where the vapour withdrawn from theflash tank is condensed, and C_(H2) and C_(Et) are the molarconcentrations in the flash tank of hydrogen and ethylene respectively.The invention assists in achieving this by minimising the concentrationof C_(lights) in the second reactor. It will be understood that “first”and “second” reactors refers to the order of polymerisation, regardlessof which polymer is made in which reactor.

Preferably the concentration of components having a molecular weightbelow 50 in the slurry entering the flash tank is controlled bycontrolling that concentration in the second reactor. Accordingly it ispreferred that the concentration in the second reactor of componentshaving a molecular weight below 50 also satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) whereC_(lights), C_(H2), and C_(Et) in this case are the concentrations ofcomponents having a molecular weight below 50, hydrogen and ethylenerespectively in the second reactor and P_(c) and T_(c) are as previouslydefined. More preferably the concentration of components having amolecular weight below 50 in the second reactor is the same as theconcentration of components having a molecular weight below 50 enteringthe flash tank.

It is generally preferred that the concentration of components having amolecular weight below 50 satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) whereC_(lights), C_(H2), C_(Et), P_(c) and T_(c) are as defined previouslyand refer either to the second reactor or the flash tank depending onthe particular embodiment of the invention.

It is preferably ensured that the concentration of components having amolecular weight below 50 in the second reactor satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)) by ensuringthe ratio of the average activity in the second LMW reactor to theaverage activity in the first HMW reactor is from 0.25 and 1.5. Averageactivity is typically higher in the first reactor (where a copolymer isusually made to obtain the HMW product) than in the second reactor(where a homopolymer is usually made to obtain the LMW product), and wehave found that as a consequence the ratio of average activities betweenthe reactors has to be controlled within these ranges in order tocontrol the concentration of light components in the second reactor.

By maintaining the preferred ratio of average activity and ethyleneconcentration ratio between the two reactors, it is possible to achievehigh overall space time yields (defined as production of polymer in kg/hper unit volume of reactor) and activities whilst still observing theC_(lights) requirements of the invention in the flash tank. The averagespace time yield in all reactors combined may be maintained at greaterthan 100 kg/m³/h, more preferably greater than 150 kg/m³/h, and mostpreferably greater than 200 kg/m³/h.

This invention is particularly applicable when the polymerisationcatalyst is a Ziegler-Natta catalyst, especially if the overallproductivity of the process is at least 10 kg polyethylene/g catalyst,preferably greater than 15 kg polyethylene/g catalyst, most preferablygreater than 20 kg polyethylene/g catalyst. If the polymerisationcatalyst is a bis-Cp metallocene catalyst, most preferably abis-tetrahydroindenyl (THI) compound, the overall productivity of theprocess in this case is preferably at least 3 kg polyethylene/gcatalyst, preferably greater than 6 kg polyethylene/g catalyst, mostpreferably greater than 15 kg polyethylene/g catalyst. If thepolymerisation catalyst is a mono-Cp metallocene catalyst, mostpreferably (t-butylamido) (tetramethyl-η⁵-cyclopentadienyl) dimethylsilanetitanium-η⁴-1.3-pentadiene, the overall productivity of theprocess in this case is preferably at least 3 kg polyethylene/gcatalyst, preferably greater than 6 kg polyethylene/g catalyst, mostpreferably greater than 15 kg polyethylene/g catalyst.

In order to achieve the above ratio of average activities, it ispreferred that the ratio of ethylene concentration in the liquid (in mol%) in the second reactor to that in the first reactor is 5 or less.Preferably the ratio of ethylene concentration in the second reactor tothat in the first reactor is 3 or less, and more preferably 2.5 or less.Most preferably both ethylene concentration ratio and average activityratio requirements are satisfied together. The ethylene concentration inthe liquid is calculated as moles of ethylene divided by moles of allliquid components.

It is preferred that the actual concentration of ethylene in the secondreactor is less than 8 mol %. However in order to ensure a satisfactorylevel of productivity, it is also preferred that the ethyleneconcentration is greater than 1.5 mol %, preferably greater than 2 mol%. The concentration of hydrogen in the second reactor is preferablyless than 5 mol %, more preferably less than 3 mol %. The ratio ofhydrogen to ethylene is preferably 0-0.5 mol/mol.

It is preferred to maintain the temperature of the first reactor between60 and 80° C., preferably less than 75° C., as this can assist inbalancing the activities between the reactors and the respective coolingcapacities.

Additives to enhance average activity may be added, preferably to theLMW reactor. Equally by-product suppressors may be added, preferably tothe LMW reactor. Additionally or alternatively, further catalyst mayalso be added to the second reactor in order to control the averageactivity balance. When operating HMW-LMW configuration it is preferredto avoid the use of an activity enhancer in the HMW reactor and inLMW-HMW configuration it can generally be avoided, however can be usedto minimise the concentration of monomers required in the HMW reactor.This reduces the downstream degassing energy requirements.

In all embodiments of the invention, one benefit of the invention isthat optimisation of reactor average activity balance, space time yieldsand cooling requirements, whilst at the same time minimising theC_(lights) concentration in the flash tank so as to avoid the need torecompress, leads to improved efficiency. This invention can enablemonomer efficiencies of less than 1.015, generally less than 1.01 andpreferably less then 1.006 to be achieved even when employing a spacetime yield of at least 100 kg/m³/h, more preferably at least 150kg/m³/h, most preferably at least 200 kg/m³/h in each reactor. By“monomer efficiency” is meant the weight ratio of ethylene+comonomerconsumed to polymer produced.

In the case where the catalyst used for the polymerisation reaction is aZiegler-Natta catalyst, it is preferred that a single activity enhancerand by-product suppressor is used in the LMW reactor. An example is ahalogenated hydrocarbon, and more particularly a chloromethane of theformula CH_(x)Cl_(4-x) where x is an integer from 1 to 3. The mostpreferred chloromethane is chloroform, CHCl₃. The amount of halogenatedhydrocarbon added is based on the amount of Ziegler-Natta catalyst, andis preferably such that the molar ratio of the halogenated hydrocarbonadded to the reactor to titanium added to the reactor is greater than0.1, preferably between 0.2 and 1. The use of a halogenated hydrocarbonis particularly desirable when used in conjunction with catalyst systemswhere it both enhances activity and suppresses the ethane formation,such as Ziegler-Natta catalysts. It is also useful in a reactorproducing low molecular weight polymer since it has the combined effectof enhancing activity and suppressing ethane formation. Ethane formationadds to the concentration of light reagents in the reactor, therebymaking it more difficult to maintain the concentration of C_(lights) inthe feed to the flash tank below the level requried by the invention.Ethane formation can be particularly significant when making lowmolecular weight polymers, particularly if hydrogen is present. Whenmaking low molecular weight polymer in the second reactor it is alsoparticularly desirable to boost the activity of the catalyst since theage of the catalyst and the high hydrogen concentration both contributeto a reduction in polymerisation activity. Halogenated hydrocarbons suchas chloroform can therefore provides a double benefit, by boostingactivity boost also minimising the concentration of C_(lights) in thesecond reactor.

A preferred type of reactor utilised for such polymerisations is a loopreactor, which is a continuous tubular construction comprising at leasttwo, for example four, vertical sections and at least two, for examplefour horizontal sections. The heat of polymerisation is typicallyremoved using indirect exchange with a cooling medium, preferably water,in jackets surrounding at least part of the tubular loop reactor. Thevolume of one loop reactor in a multiple reactor system can vary but istypically in the range 10-200 m³. It is preferred that thepolymerisation reactor utilised in the present invention is a loopreactor.

Typical pressures employed in the loop reactor are between 0.1-10 MPa g,preferably between 3 to 5 MPa g.

The process according to the invention applies to the preparation ofcompositions containing ethylene homopolymers and copolymers. Ethylenecopolymers typically comprise one or more alpha-olefins in a variableamount which can reach 12% by weight, preferably from 0.5 to 6% byweight, for example approximately 1% by weight.

The alpha mono-olefin monomers generally employed in such reactions areone or more 1-olefins having up to 8 carbon atoms per molecule and nobranching nearer the double bond than the 4-position. Typical examplesinclude ethylene, propylene, butene-1, pentene-1, hexene-1 and octene-1,and mixtures such as ethylene and butene-1 or ethylene and hexene-1.Butene-1, pentene-1 and hexene-1 are particularly preferred comonomersfor ethylene copolymerisation.

In one embodiment of the invention, the polymer is a polyethylene resinhaving a density of greater than 940 kg/m³ and an HLMI of from 1 to 100g/10 min, and comprising from 35 to 60 wt % of a first polyethylenefraction of high molecular weight and from 40 to 65 wt % of a secondpolyethylene fraction of low molecular weight, the first polyethylenefraction comprising a linear low density polyethylene having a densityof up to 935 kg/m³ and an HLMI of less than 1 g/10 min, and the secondpolyethylene fraction comprising a high density polyethylene having adensity of at least 960 kg/m³, preferably at least 965 kg/m³, and an MI₂of greater than 100 g/10 min, and the polyethylene resin.

Typical diluents for the suspensions in each reactor includehydrocarbons having 2 to 12, preferably 3 to 8, carbon atoms permolecule, for example linear alkanes such as propane, n-butane, n-hexaneand n-heptane, or branched alkanes such as isobutane, isopentane,isooctane and 2,2-dimethylpropane, or cycloalkanes such as cyclopentaneand cyclohexane or their mixtures. In the case of ethylenepolymerization, the diluent is generally inert with respect to thecatalyst, cocatalyst and polymer produced (such as liquid aliphatic,cycloaliphatic and aromatic hydrocarbons), at a temperature such that atleast 50% (preferably at least 70%) of the polymer formed is insolubletherein. Isobutane is particularly preferred as the diluent.

The operating conditions can also be such that the monomers act as thediluent as is the case in so called bulk polymerisation processes. Theslurry concentration limits in volume percent have been found to be ableto be applied independently of molecular weight of the diluent andwhether the diluent is inert or reactive, liquid or supercritical.Propylene monomer is particularly preferred as the diluent for propylenepolymerisation

Methods of molecular weight regulation are known in the art. When usingZiegler-Natta, metallocene and tridentate late transition metal typecatalysts, hydrogen is preferably used, a higher hydrogen pressureresulting in a lower average molecular weight. When using chromium typecatalysts, polymerization temperature is preferably used to regulatemolecular weight.

In commercial plants, the particulate polymer is separated from thediluent in a manner such that the diluent is not exposed tocontamination so as to permit recycle of the diluent to thepolymerization zone with minimal if any purification. Separating theparticulate polymer produced by the process of the present inventionfrom the diluent typically can be by any method known in the art forexample it can involve either (i) the use of discontinuous verticalsettling legs such that the flow of slurry across the opening thereofprovides a zone where the polymer particles can settle to some extentfrom the diluent or (ii) continuous product withdrawal via a single ormultiple withdrawal ports, the location of which can be anywhere on theloop reactor but is preferably adjacent to the downstream end of ahorizontal section of the loop. The operation of large diameter reactorswith high solids concentrations in the slurry minimises the quantity ofthe principal diluent withdrawn from the polymerisation loop. Use ofconcentrating devices on the withdrawn polymer slurry, preferablyhydrocylones (single or in the case of multiple hydrocyclones inparallel or series), further enhances the recovery of diluent in anenergy efficient manner since significant pressure reduction andvaporisation of recovered diluent is avoided. Increasing theconcentration of easily condensable components, for example throughaddition of fresh or recycle diluent, upstream of the hydrocyclone is afurther means of enhancing the operating window of the final reactor andreducing the concentration of monomer depressurised to the mediumpressure flash tank.

Where the final reactor of the multiple reactor system is a loopreactor, the withdrawn, and preferably concentrated, polymer slurry isdepressurised, and optionally heated, prior to introduction into aprimary flash vessel. The stream is preferably heated afterdepressurisation. As a consequence of the invention, the diluent and anymonomer vapours recovered in the primary flash vessel can be condensedwithout recompression. They are typically then recycled to thepolymerization process. Typically the pressure in the primary flashvessel is 0.5-2.5 MPa g, preferably 0.5-1.5 MPa g. The solids recoveredfrom the primary flash vessel are usually passed to a secondary flashvessel to remove residual volatiles.

The process according to the invention is relevant to all olefinpolymerisation catalyst systems, particularly those chosen from theZiegler-type catalysts, in particular those derived from titanium,zirconium or vanadium and from thermally activated silica or inorganicsupported chromium oxide catalysts and from metallocene-type catalysts,metallocene being a cyclopentadienyl derivative of a transition metal,in particular of titanium or zirconium.

Non-limiting examples of Ziegler-type catalysts are the compoundscomprising a transition metal chosen from groups IIIB, IVB, VB or VIB ofthe periodic table, magnesium and a halogen obtained by mixing amagnesium compound with a compound of the transition metal and ahalogenated compound. The halogen can optionally form an integral partof the magnesium compound or of the transition metal compound.

Metallocene-type catalysts may be metallocenes activated by either analumoxane or by an ionising agent as described, for example, in EP500944A (Mitsui Toatsu Chemicals).

Ziegler-type catalysts are most preferred. Among these, particularexamples include at least one transition metal chosen from groups IIILB,IVB, VB and VIB, magnesium and at least one halogen. Good results areobtained with those comprising:

from 10 to 30% by weight of transition metal, preferably from 15 to 20%by weight,

from 20 to 60% by weight of halogen, preferably from 30 to 50% by weight

from 0.5 to 20% by weight of magnesium, usually from 1 to 10% by weight,

from 0.1 to 10% by weight of aluminium, generally from 0.5 to 5% byweight,

the balance generally consists of elements arising from the productsused for their manufacture, such as carbon, hydrogen and oxygen. Thetransition metal and the halogen are preferably titanium and chlorine.Most preferred catalysts have the following composition:

Transition metal from 8 to 20% by weight

Magnesium content from 3 to 15% by weight

Chlorine content from 40 to 70% by weight

Aluminum content less than 5% by weight

Residual organic content less than 40% by weight

Polymerisations, particularly Ziegler catalysed ones, are typicallycarried out in the presence of a cocatalyst. It is possible to use anycocatalyst known in the art, especially compounds comprising at leastone aluminium-carbon chemical bond, such as optionally halogenatedorganoaluminium compounds, which can comprise oxygen or an element fromgroup I of the periodic table, and aluminoxanes. Particular exampleswould be organoaluminium compounds, of trialkylaluminiums such astriethylaluminium, trialkenylaluminiums such as triisopropenylaluminium,aluminium mono- and dialkoxides such as diethylaluminium ethoxide, mono-and dihalogenated alkylaluminiums such as diethylaluminium chloride,alkylaluminium mono- and dihydrides such as dibutylaluminium hydride andorganoaluminium compounds comprising lithium such as LiAl(C₂H₅)₄.Organoaluminium compounds, especially those which are not halogenated,are well suited. Triethylaluminium and triisobutylaluminium areespecially advantageous.

In one particular embodiment of the invention, the catalyst employed inthe process is a Ziegler-Natta catalyst, the weight ratio of LMW to BMWpolymer is from 40:60 to 60:40, and the space time yield (defined asproduction of polymer in kg/h per unit volume of reactor) is at least150, preferably at least 200, most preferably at least 250.

The chromium-based catalyst is preferred to comprise a supportedchromium oxide catalyst having a titania-containing support, for examplea composite silica and titania support. A particularly preferredchromium-based catalyst may comprise from 0.5 to 5 wt % chromium,preferably around 1 wt % chromium, such as 0.9 wt % chromium based onthe weight of the chromium-containing catalyst. The support comprises atleast 2 wt % titanium, preferably around 2 to 3 wt % titanium, morepreferably around 2.3 wt % titanium based on the weight of the chromiumcontaining catalyst. The chromium-based catalyst may have a specificsurface area of from 200 to 700 m²/g, preferably from 400 to 550 m²/gand a volume porosity of greater than 2 cc/g preferably from 2 to 3cc/g. Chromium-based catalysts may be used in conjunction withactivators such organometallic compounds of aluminium or of boron.Preferred are organoboron compounds such as trialkylborons in which thealkyl chains comprise up to 20 carbon atoms. Triethylboron isparticularly preferred.

If the catalyst employed is a metallocene catalyst, it preferablycomprises a bis-tetrahydroindenyl (THD) compound. Preferably thecatalyst system comprises (a) a metallocene catalyst componentcomprising a bis-tetrahydroindenyl compound of the general formula(IndH₄)₂R″MQ₂ in which each IndH₄ is the same or different and istetrahydroindenyl or substituted tetrahydroindenyl, R″ is a bridge whichcomprises a C₁-C₄ alkylene radical, a dialkyl germanium or silicon orsiloxane, or an alkyl phosphine or amine radical, which bridge issubstituted or unsubstituted, M is a Group IV metal or vanadium and eachQ is hydrocarbyl having 1 to 20 carbon atoms or halogen; and (b) acocatalyst which activates the catalyst component. Eachbis-tetrahydroindenyl compound may be substituted in the same way ordifferently from one another at one or more positions in thecyclopentadienyl ring, the cyclohexenyl ring and the ethylene bridge.Each substituent group may be independently chosen from those of formulaXR_(v) in which X is chosen from group IVB, oxygen and nitrogen and eachR is the same or different and chosen from hydrogen or hydrocarbyl offrom 1 to 20 carbon atoms and v+1 is the valence of X. X is preferablyC. If the cyclopentadienyl ring is substituted, its substituent groupsmust not be so bulky as to affect coordination of the olefin monomer tothe metal M. Substituents on the cyclopentadienyl ring preferably have Ras hydrogen or CH₃. More preferably, at least one and most preferablyboth cyclopentadienyl rings are unsubstituted. In a particularlypreferred embodiment, both indenyls are unsubstituted. R″ is preferablyan ethylene bridge which is substituted or unsubstituted. The metal M ispreferably zirconium, hafnium or titanium, most preferably zirconium.Each Q is the same or different and may be a hydrocarbyl or hydrocarboxyradical having 1-20 carbon atoms or a halogen. Suitable hydrocarbylsinclude aryl, alkyl, alkenyl, alkylaryl or aryl alkyl. Each Q ispreferably halogen. Ethylene bis(4,5,6,7-tetrahydro-1-indenyl) zirconiumdichloride is a particularly preferred bis-tetrahydroindenyl compound.

Silica supported chromium catalysts are typically subjected to aninitial activation step in air at an elevated activation temperature.The activation temperature preferably ranges from 500 to 850° C., morepreferably 600 to 750° C.

In the process of the invention, the first reactor of the series issupplied with catalyst and the cocatalyst in addition to the diluent andmonomer, and each subsequent reactor is supplied with, at least,monomer, in particular ethylene and with the slurry arising from apreceding reactor of the series, this mixture comprising the catalyst,the cocatalyst and a mixture of the polymers produced in a precedingreactor of the series. It is optionally possible to supply a secondreactor and/or, if appropriate, at least one of the following reactorswith fresh catalyst and/or cocatalyst. However, it is preferable tointroduce the catalyst and the cocatalyst exclusively into a firstreactor.

1. Process for producing a multimodal polyethylene in at least tworeactors connected in series, in which 20-80 wt % of a high molecularweight (HMW) polymer is made in suspension in a first reactor and 20-80wt % of a low molecular weight (LMW) polymer is made in suspension in asecond reactor, wherein the ratio of the average activity in the LMWreactor to the average activity in the HMW reactor is from 0.25 to 1.5,where average activity in each reactor is defined as the rate ofpolyethylene produced in the reactor (kgPE/hr)/[ethylene concentrationin the reactor (mol %)×residence time in the reactor (hours)×feed rateof catalyst into the reactor (g/hr)], residence time being defined asthe mass of the polymer in the reactor (kg)/the output rate of polymerfrom the reactor (kg/hr), and wherein the volume of the second reactoris at least 10% greater than the volume of the first reactor, and theratio of length to diameter of the first reactor, L/D(1), is greaterthan that of the second reactor, L/D(2).
 2. Process according to claim1, wherein the volume of the second reactor is at least 30% greater thanthe volume of the first reactor.
 3. Process according to claim 1,wherein the volume of the second reactor is at least 50% greater thanthe volume of the first reactor.
 4. Process according to claim 1,wherein the ratio of length to diameter of the first reactor, L/D(1), isat least 20% greater than that of the second reactor, L/D(2).
 5. Processaccording to claim 1, wherein the solids concentration in the second LMWreactor, defined as the mass of polymer divided by the total mass ofslurry, is at least 35 wt %, and/or the ratio of solids concentration inthe first reactor to that in the second reactor is maintained at lessthan 1.0.
 6. Process according to claim 5, wherein the solidsconcentration in the second LMW reactor is between 45 wt % and 60 wt %.7. Process according to claim 5, wherein the ratio of solidsconcentration in the first reactor to that in the second reactor ismaintained between 0.6 and 0.8.
 8. Process for producing a multimodalpolyethylene in at least two reactors connected in series, in which20-80 wt % of a high molecular weight (HMW) polymer is made insuspension in a first reactor and 20-80 wt % of a low molecular weight(LMW) polymer is made in suspension in a second reactor, in which thesolids concentration in the second LMW reactor, defined as the mass ofpolymer divided by the total mass of slurry, is at least 35 wt %, and/orthe ratio of solids concentration in the first reactor to that in thesecond reactor is maintained at less than 1.0, and the volume of thesecond reactor is at least 10%, and the ratio of length to diameter ofthe first reactor, L/D(1), is greater than that of the second reactor,L/D(2).
 9. Process according to claim 8, wherein the solidsconcentration in the second LMW reactor is between 45 wt % and 60 wt %.10. Process according to claim 8, wherein the volume of the secondreactor is at least 30% greater than the volume of the first reactor.11. Process according to claim 8, wherein the volume of the secondreactor is at least 50% greater than the volume of the first reactor.12. Process according to claim 8, wherein the ratio of length todiameter of the first reactor, L/D(1), is at least 20% greater than thatof the second reactor, L/D(2).
 13. Process according to claim 8, whereinthe ratio of solids concentration in the first reactor to that in thesecond reactor is maintained between 0.6 and 0.8.
 14. Process accordingto claim 8, wherein the ratio of the average activity in the LMW reactorto the average activity in the HMW reactor is from 0.25 and 1.5, whereaverage activity in each reactor is defined as the rate of polyethyleneproduced in the reactor (kgPE/hr)/[ethylene concentration in the reactor(mol %)×residence time in the reactor (hours)×feed rate of catalyst intothe reactor (g/hr)], residence time being defined as the mass of thepolymer in the reactor (kg)/the output rate of polymer from the reactor(kg/hr).
 15. Process according to claim 1 or 8, wherein the solidsconcentration in the first HMW reactor is between 20 wt % and 50 wt %.16. Process according to claim 15, wherein the solids concentration inthe first HMW reactor is between 25 wt % and 35 wt %.
 17. Processaccording to claim 1 or 8, wherein the ratio L/D(1) to L/D(2) is greaterthan 1.5.
 18. Process according to claim 1 or 8, wherein the ratioL/D(1) to L/D(2) is greater than
 2. 19. Process according to claim 1 or8, wherein the ratio of length to diameter (L/D) of the first HMWreactor is greater than
 500. 20. Process according to claim 19, whereinthe ratio of length to diameter (L/D) of the first HMW reactor isgreater than
 800. 21. Process according to claim 1 or 8, wherein theratio of length to diameter (L/D) of the second LMW reactor is greaterthan
 200. 22. Process according to claim 1 or 8, wherein the averagespace time yield (defined as production of polymer in kg/h per unitvolume of reactor) in all reactors combined is greater than 100 kg/m³/h.23. Process according to claim 1 or 8, wherein the catalyst employed inthe process is a Ziegler-Natta catalyst, the weight ratio of LMW to HMWpolymer is from 40:60 to 60:40, and the space time yield (defined asproduction of polymer in kg/h per unit volume of reactor) is at least150.
 24. Process according to claim 1 or 8, wherein the ratio ofethylene concentration in the liquid phase (in mol %) in the secondreactor to that in the first reactor is 5 or less.
 25. Process accordingto claim 1 or 8, wherein the concentration of ethylene in the secondreactor is less than 8 mol %.
 26. Process according to claim 1 or 8,wherein the temperature of the first reactor is maintained between 60and 80° C.
 27. Process according to claim 1 or 8, wherein a slurrycontaining the multimodal polyethylene is transferred from the second ofthe two reactors to a flash tank operating at a pressure and temperaturesuch that at least 50 mol % of the liquid component of the slurry iswithdrawn from the flash tank as a vapour.
 28. Process according toclaim 27, wherein the concentration in the second reactor of componentshaving a molecular weight below 50 also satisfies the equationC_(lights)<7+0.07(40−T_(c))+4.4(P_(c)−0.8)−7(C_(H2)/C_(Et)), whereC_(lights), C_(H2), and C_(Et) in this case are the concentrations ofcomponents having a molecular weight below 50, hydrogen and ethylenerespectively in the second reactor, T_(c) is the condensationtemperature (° C.) of said vapour, and P_(c) is the pressure (MPa g) atthe location where the vapour withdrawn from the flash tank iscondensed.